Apparatus and process for light olefin recovery

ABSTRACT

The present invention relates to a process and apparatus for the production of light olefins comprising olefins having from 2 to 3 carbon atoms per molecule from a feedstock containing heavier olefins. An intermediate cut from a fractionation column is used as olefinic feed to an olefin cracking process preferably after undergoing selective hydrogenation of diolefins. In one embodiment, a liquid side draw from a fractionation column is selectively hydrogenated and then returned to the fractionation column from which a vapor side draw containing olefins is cracked in the olefin cracking reactor.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a Division of copending application Ser. No.10/882,531 filed Jun. 30, 2004, the contents of which are herebyincorporated by reference in its entirety.

FIELD OF THE INVENTION

This invention relates to an apparatus and process for the recovery oflight olefins produced by cracking heavier olefins. This invention moreparticularly refers to taking an intermediate cut from fractionationcolumn as feed to an olefin cracking reactor.

BACKGROUND OF THE INVENTION

Light olefins, ethylene and propylene, serve as feeds for plasticsproduction and for the production of petrochemicals which serve as feedsfor plastics production. The demand for light olefins has been steadilyincreasing and will continue to increase dramatically. Light olefinshave traditionally been produced through the process of steam orcatalytic cracking. Paraffin dehydrogenation is an alternative source oflight olefins. However, the demand for light olefins is outstripping thecapacity of traditional sources of light olefins.

The search for alternative materials for light olefin production has ledto the use of oxygenates such as alcohols and, more particularly, to theuse of methanol, ethanol and higher alcohols or their derivatives. Thesealcohols may be produced by fermentation or from synthesis gas.Synthesis gas can be produced from natural gas, petroleum liquids andfrom carbonaceous materials including coal, recycled plastics, municipalwastes, or any organic material. Thus, alcohol and alcohol derivativesmay provide non-petroleum based routes for the production of olefins andother hydrocarbons. Methanol, in particular, is useful in this processwhich is referred to herein as the methanol-to-olefin (MTO) process.

Molecular sieve catalysts such as the microporous crystalline zeoliteand non-zeolitic catalysts, particularly silicoaluminophosphates (SAPO),are known to promote the conversion of oxygenates to hydrocarbonmixtures. Numerous patents describe this process for various types ofthese catalysts: U.S. Pat. No. 3,928,483; U.S. Pat. No. 4,025,575; U.S.Pat. No. 4,252,479; U.S. Pat. No. 4,496,786; U.S. Pat. No. 4,547,616;U.S. Pat. No. 4,677,243; U.S. Pat. No. 4,843,183; U.S. Pat. No.4,499,314; U.S. Pat. No. 4,447,669; U.S. Pat. No. 5,095,163; U.S. Pat.No. 5,191,141; U.S. Pat. No. 5,126,308; U.S. Pat. No. 4,973,792; andU.S. Pat. No. 4,861,938.

The MTO process may be generally conducted in the presence of one ormore diluents which may be present in the oxygenate-containing feed inan amount between about 1 and about 99 mol-%, based on the total numberof moles of all feed and diluent components fed to the reaction zone.Diluents include, but are not limited to, helium, argon, nitrogen,carbon monoxide, carbon dioxide, hydrogen, water, paraffins,hydrocarbons (such as methane and the like), aromatic compounds, ormixtures thereof. U.S. Pat. No. 4,861,938 and U.S. Pat. No. 4,677,242particularly emphasize the use of a diluent combined with the feed tothe reaction zone to maintain sufficient catalyst selectivity toward theproduction of light olefin products, particularly ethylene.

Generally, the product ratio of ethylene to propylene on a carbon basisvaries from about 0.1 to about 10 and, more typically, varies from about0.8 to about 2.5. Ethylene and propylene are particularly desirableolefins but it has been found that their yields are reduced by theproduction of medium-weight hydrocarbons such as C₄ to C₈ olefins, aswell as some heavier components. U.S. Pat. No. 5,914,433 proposescracking medium-weight olefins over a catalyst in vapor phase toincrease overall yield of light olefins.

A portion of the medium-weight olefin stream, when cracked, will beconverted to paraffinic compounds such as methane, ethane, propane, andheavier hydrocarbons. Unless at least a portion of these compounds areremoved, they will build up in the system and reduce the overallefficiency of the process. Therefore, a drag stream comprising C₄ andheavier hydrocarbons is removed from the process and used for plantfuel, blended into other hydrocarbon products such as motor gasoline orused as feed to a gasoline alkylation process.

In order to maximize the production of light olefins and to minimize theproduction of methane produced from cracking the medium-weight olefinstream, diolefins and acetylenes should be minimized in the feed to theolefin cracking zone. Diolefin and acetylene conversion to monoolefinhydrocarbons may be accomplished with a conventional selectivehydrogenation process such as disclosed in U.S. Pat. No. 4,695,560.

In a light olefins recovery flow scheme, U.S. Pat. No. 6,486,369discloses a single selective hydrogenation converter for treating methylacetylene and propadiene in a feed to a deethanizer column. WO2004/009519 A1 discloses a fractionation scheme for recovering lightolefins from a medium-weight olefin containing stream.

SUMMARY OF THE INVENTION

In the present invention, an intermediate cut is taken from a productfractionation column and delivered to the olefin cracking reactor.Additionally, the intermediate cut can be taken below a tray to withdrawa vapor stream from the fractionation column for feed to the olefincracking reactor thereby obviating the need to vaporize the feed to theolefin cracking reactor. Moreover, liquid olefinic feed is mixed with aliquid side cut taken from a fractionation tray and hydrogen andselectively hydrogenated in a reactor to reduce concentrations ofdiolefins and acetylenes. Hydrogenated feed returned to a lower sectionof the column ascends in the column until withdrawn as vapor to becracked in the olefin cracking reactor. A C₃ and lighter product streamis withdrawn from the overhead of the fractionation column.

An object of the present invention is to simplify the recovery sectionof an olefin cracking process by incorporating an intermediate cut of afractionation column as feed to the olefin cracking reactor.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic process flow diagram illustrating the process andapparatus of the present invention.

FIG. 2 is a schematic process flow diagram illustrating an alternativeprocess and apparatus of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The invention comprises a process and apparatus for the catalyticconversion of an olefinic feed stream containing intermediate-weight C₄to C₇ olefins to a cracked product stream containing light olefins, C₂to C₃ olefins. The olefinic feed stream may be obtained from a mid-cut,C₄ to C₈ range, of a fluid catalytic cracking (FCC) product stream orfrom a C₄ to C₆ stream of a steam cracker furnace which contain abundantolefinic species in the desired intermediate weight range. Moreover, aC₄ ⁺ product from an MTO unit, which converts oxygenates over asilicoaluminophosphate molecular sieve catalyst to light olefins asdescribed in U.S. Pat. No. 5,914,433, would also serve well forupgrading overall process selectivity to light olefins. Increased yieldof the light-weight olefinic products for all of the processes isprovided by sending C₄ to C₈ medium-weight olefins to an olefin crackingreactor. Depending upon operating conditions, the medium-weight olefinsmay be C₄ to C₅-C₇ olefins.

Catalysts suitable for olefin cracking comprise a crystalline silicateof the MFI family which may be a zeolite, a silicalite or any othersilicate in that family or the MEL family which may be a zeolite or anyother silicate in that family. Examples of MFI silicates are ZSM-5 andSilicalite. An example of an MEL zeolite is ZSM-11 which is known in theart. Other examples are Boralite D and silicalite-2 as described by theInternational Zeolite Association (ATLAS OF ZEOLITE STRUCTURE TYPES,1987, Butterworths). The preferred crystalline silicates have pores orchannels defined by ten oxygen rings and a high silicon/aluminum atomicratio.

The crystalline silicate catalyst has structural and chemical propertiesand is employed under particular reaction conditions whereby thecatalytic cracking of the C₄ to C₇ olefins readily proceeds. Differentreaction pathways can occur on the catalyst. Suitable olefin crackingprocess conditions include an inlet temperature of around 400° to 600°C., preferably from 520° to 600° C., yet more preferably 540° to 580°C., and an olefin partial pressure of from 10 to 202 kPa absolute (1.5to 29 psia), preferably from 50 to 152 kPa absolute (7 to 22 psia).Furthermore, such isomerization tends to reach a thermodynamicequilibrium. Olefinic catalytic cracking may be understood to comprise aprocess yielding shorter molecules via bond breakage.

A crystalline silicate catalyst possessing a high silicon/aluminum ratiocan achieve a stable olefin conversion with a high propylene yield on anolefin basis of from 20 to 50 wt-% with the olefinic feedstocks of thepresent invention. The MFI catalyst having a high silicon/aluminumatomic ratio for use in the catalytic olefin cracking process of thepresent invention may be manufactured by removing aluminum from acommercially available crystalline silicate. A typical commerciallyavailable Silicalite has a silicon/aluminum atomic ratio of around 120.The commercially available MFI crystalline silicate may be modified by asteaming process which reduces the tetrahedral aluminum in thecrystalline silicate framework and converts the aluminum atoms intooctahedral aluminum in the form of amorphous alumina. Although in thesteaming step aluminum atoms are chemically removed from the crystallinesilicate framework structure to form alumina particles, those particlescause partial obstruction of the pores or channels in the framework.This inhibits the olefin cracking process. Accordingly, following thesteaming step, the crystalline silicate is subjected to an extractionstep wherein amorphous alumina is removed from the pores and themicropore volume is, at least partially, recovered. The physicalremoval, by a leaching step, of the amorphous alumina from the pores bythe formation of a water-soluble aluminum complex yields the overalleffect of de-alumination of the MFI crystalline silicate. In this way byremoving aluminum from the MFI crystalline silicate framework and thenremoving alumina formed therefrom from the pores, the process aims atachieving a substantially homogeneous de-alumination throughout thewhole pore surfaces of the catalyst. This reduces the acidity of thecatalyst and thereby reduces the occurrence of hydrogen transferreactions in the cracking process. The reduction of acidity ideallyoccurs substantially homogeneously throughout the pores defined in thecrystalline silicate framework. This is because in the olefin-crackingprocess hydrocarbon species can enter deeply into the pores.Accordingly, the reduction of acidity and thus the reduction in hydrogentransfer reactions which would reduce the stability of the MFI catalystare pursued throughout the whole pore structure in the framework. Theframework silicon/aluminum ratio may be increased by this process to avalue of at least about 180, preferably from about 180 to 1000, morepreferably at least 200, yet more preferably at least 300 and mostpreferably around 480.

The MEL or MFI crystalline silicate catalyst may be mixed with a binder,preferably an inorganic binder, and shaped to a desired shape, e.g.extruded pellets. The binder is selected so as to be resistant to thetemperature and other conditions employed in the catalyst manufacturingprocess and in the subsequent catalytic cracking process for theolefins. The binder is an inorganic material selected from clays,silica, metal oxides such as ZrO₂ and/or metals, or gels includingmixtures of silica and metal oxides. The binder is preferablyalumina-free, although aluminum in certain chemical compounds as inAlPO₄ may be used as the latter are quite inert and not acidic innature. If the binder which is used in conjunction with the crystallinesilicate is itself catalytically active, this may alter the conversionand/or the selectivity of the catalyst. Inactive materials for thebinder may suitably serve as diluents to control the amount ofconversion, so that products can be obtained economically and orderlywithout employing other means for controlling the reaction rate. It isdesirable to provide a catalyst having a good crush strength to preventthe catalyst from breaking down into powder-like materials during use.Such clay or oxide binders have been employed normally for the purposeof improving the crush strength of the catalyst. A particularlypreferred binder for the catalyst of the present invention comprisessilica or AlPO₄.

The relative proportions of the finely divided crystalline silicatematerial and the inorganic oxide matrix of the binder can vary widely.Typically, the binder content ranges from 5 to 95% by weight, moretypically from 20 to 50% by weight, based on the weight of the compositecatalyst. Such a mixture of crystalline silicate and an inorganic oxidebinder is referred to as a formulated crystalline silicate.

In mixing the catalyst with a binder, the catalyst may be formulatedinto pellets, spheres, extruded into other shapes, or formed into aspray-dried powder. In the catalytic cracking process, the processconditions are selected in order to provide high selectivity towardspropylene or ethylene, as desired, a stable olefin conversion over time,and a stable olefinic product distribution in the effluent. Suchobjectives are favored by the use of a low acid density in the catalyst(i.e. a high Si/Al atomic ratio) in conjunction with a low pressure, ahigh inlet temperature and a short contact time, all of which processparameters are interrelated and provide an overall cumulative effect.

The process conditions are selected to disfavor hydrogen transferreactions leading to the formation of paraffins, aromatics and cokeprecursors. The process operating conditions thus employ a high spacevelocity, a low pressure and a high reaction temperature. The LHSVranges from 5 to 30 hr⁻¹, preferably from 10 to 30 hr⁻¹. The olefinpartial pressure ranges from 10 to 202 kPa absolute (1.5 to 29 psia),preferably from 50 to 152 kPa absolute (7 to 22 psia). A particularlypreferred olefin partial pressure is atmospheric pressure. Thehydrocarbon feedstocks are preferably fed at a total inlet pressuresufficient to convey the feedstocks through the reactor. The hydrocarbonfeedstocks may be fed undiluted or diluted in an inert gas, e.g.nitrogen or steam. The total absolute pressure in the reactor rangesfrom 30 to 1013 kPa absolute (4 to 147 psia) and is preferablyatmospheric. The use of a low olefin partial pressure, for exampleatmospheric pressure, tends to lower the incidence of hydrogen transferreactions in the cracking process, which in turn reduces the potentialfor coke formation which tends to reduce catalyst stability. Thecracking of the olefins is preferably performed at an inlet temperatureof the feedstock of from 400° to 650° C., more preferably from 450° to600° C., yet more preferably from 540° to 590° C., typically around 560°to 585° C.

In order to maximize the amount of ethylene and propylene and tominimize the production of methane produced from the butylene andheavier stream, it is desired to minimize the presence of diolefins andacetylenes in the feed. Diolefin conversion to monoolefin hydrocarbonsmay be accomplished with a conventional selective hydrogenation catalystwhich comprises an alumina support material preferably having a totalsurface area greater than 150 m²/g, with most of the total pore volumeof the catalyst provided by pores with average diameters of greater than600 angstroms, and containing surface deposits of about 1.0 to 25.0 wt-%nickel and about 0.1 to 1.0 wt. % sulfur such as disclosed in U.S. Pat.No. 4,695,560. Spheres having a diameter between about 0.4 and 6.4 mm (1/64 and ¼ inch) can be made by oil dropping a gelled alumina sol. Thealumina sol may be formed by digesting aluminum metal with an aqueoussolution of approximately 12 wt-% hydrogen chloride to produce analuminum chloride sol. The nickel component may be added to the catalystduring the sphere formation or by immersing calcined alumina spheres ina aqueous solution of a nickel compound followed by drying, calcining,purging and reducing. The nickel containing alumina spheres may then besulfided.

The selective hydrogenation processes is normally performed atrelatively mild hydrogenation conditions. These conditions will normallyresult in the hydrocarbons being present as liquid phase materials. Thereactants will normally be maintained under the minimum pressuresufficient to maintain the reactants as liquid phase hydrocarbons whichallows the hydrogen to dissolve into the hydrocarbonaceous hydrogenationfeed. A broad range of suitable operating pressures therefore extendsfrom about 276 to 5516 kPa gauge (40 to about 800 psig), with a pressurebetween about 345 and 2069 kPa gauge (50 and 300 psig) being preferred.A relatively moderate temperature between about 25° C. and about 350° C.(77° to 662° F.) should be employed. Preferably, the temperature of thehydrogenation zone is maintained between about 30° and about 200° C.(122° and 392° F.). The liquid hourly space velocity of the reactantsthrough the selective hydrogenation catalyst should be above 1.0 hr⁻¹.Preferably, it is above 5.0 and more preferably it is between 5.0 and35.0 hr⁻¹. The ratio of hydrogen to diolefinic hydrocarbons maintainedwithin the selective hydrogenation zone is an important variable. Theamount of hydrogen required to achieve a certain conversion is believeddependent upon both reactor temperature and the molecular weight of thefeed hydrocarbons. To avoid the undesired saturation of a significantamount monoolefinic hydrocarbons, there should be less than 2.0 timesthe stoichiometric amount of hydrogen required for the selectivehydrogenation of the diolefinic hydrocarbons which are present in theliquid phase process stream to monoolefinic hydrocarbons. Preferably,the mole ratio of hydrogen to diolefinic hydrocarbons in the materialentering the bed of selective hydrogenation catalyst is maintainedbetween 1:1 and 1.8:1. In some instances, it may be desirable to operatewith a less than stoichiometrically required amount of hydrogen, withmole ratios down to 0.75:1 being acceptable. The optimum set ofconditions will of course vary depending on such factors as thecomposition of the feed stream and the degree of saturation ofdiolefinic hydrocarbons which is desired.

The hydrogenation reactor is preferably a cylindrical fixed bed ofcatalyst through which the reactants move in a vertical direction. It ispreferred that the reactants flow upward through the reactor as thisprovides good mixing. The hydrogenation catalyst may be present withinthe reactor as pellets, spheres, extrudates, irregular shaped granules,etc. To employ the hydrogenation catalyst, the reactants would bepreferably brought to the desired inlet temperature of the reactionzone, admixed with hydrogen and then passed into and through thereactor. Alternatively, the reactants may be admixed with the desiredamount of hydrogen and then heated to the desired inlet temperature. Ineither case, the effluent of the hydrogenation reactor may be passedinto a hydrogen recovery facility for the removal of residual hydrogenbefore proceeding further in the process. Hydrogen may be removed byflashing the hydrogenation effluent stream to a lower pressure or bypassing the effluent stream into a stripping column. Otherwise, noresidual hydrogen recovery may be necessary if the residual hydrogenconcentration in the hydrogenation effluent is acceptable. The effluentfrom the selective hydrogenation reactor will preferably have less than100 ppm of diolefins. The selective hydrogenation reactor may be omittedif the concentration of diolefins is already below 100 ppm.

A portion of the olefin cracking feed, when cracked, will be convertedto paraffinic compounds such as methane, ethane, propane, and heavierhydrocarbons. Hydrogenation may have the same effect. Unless at least aportion of these paraffinic compounds is removed, they will build up inthe system. Therefore, a drag stream comprising C₄ and heavierhydrocarbons is removed from the process and used for plant fuel orblended into other hydrocarbon products such as motor gasoline.Additionally, paraffins can be returned to the FCC unit or to the steamcracker for further cracking.

A depropanizer fractionation column may be used for light olefinrecovery in the overhead. Further downstream processing may be necessaryto separate paraffins from olefins and to separately recover propyleneand ethylene. The depropanizer column may be run at a pressure of 800 to2100 kPa absolute (116 to 305 psia). Hence, a compressor may be requiredto pressurize the cracked olefin effluent before entering thedepropanizer. A debutanizer fractionation column may used to collect andredirect a portion of the depropanizer bottoms stream for processrecycle. The debutanizer column may be run at pressure similar to thedepropanizer column but is typically less.

DETAILED DESCRIPTION OF THE DRAWINGS

The following description of the present process is made with referenceto the drawing. In the interest of simplifying the description of theinvention, the process system in the drawing does not contain theseveral conduits, valves, heat exchangers, and the like which, in actualpractice, would be provided in accordance with routine skill in the artto enable the process to be carried out on a continuous basis.

FIG. 1 illustrates a first embodiment of the present invention whichutilizes two fractionation columns. A fresh olefinic feed comprising 20to 50 wt-% intermediate-weight C₄ to C₈ hydrocarbons with olefins entersthe process in a line 10. A line 12 carrying a liquid process side cutcontaining intermediate-weight C₅ to C₇ hydrocarbons with olefins from adebutanizer column 16 mixes with the line 10 to provide an enriched feedin a line 18. The enriched feed in the line 18 mixes with a processportion in a line 20 of a C₄ ⁻ stream in a line 22 from the overhead ofthe debutanizer column 16 and then mixes with recycled hydrogenationeffluent in a line 24 from a hydrogenation reactor 26 to provide anolefin hydrogenation feed in a line 28. The olefin hydrogenation feedincludes diolefins and acetylenes which would convert to coke in aheater 32 preceding an olefin cracking reactor 34 and can crack tomethane in the olefin cracking reactor 34. Hence, the olefinhydrogenation feed in the line 28 is admixed with hydrogen from a line36 and heated in a heat exchanger 38 before the hydrogen containingolefin hydrogenation feed in a line 40 enters the selectivehydrogenation reactor 26. The olefin hydrogenation feed is in liquidphase and the hydrogen is dissolved into the liquid hydrocarbonaceoushydrogenation feed. The liquid olefin hydrogenation feed enters theselective hydrogenation reactor 26 and contacts a preferably fixed bedof hydrogenation catalyst. The hydrogenation reactor 26 may beconfigured for up flow or radial flow, even though down flow is shown inFIG. 1. Under selective hydrogenation conditions, diolefins andacetylenes in the olefin hydrogenation feed are converted to monoolefinswithout substantial monoolefin saturation. Hydrogenation effluent with asmaller concentration of diolefins than in the olefin hydrogenation feedin the line 40 exits the hydrogenation reactor 26 in a line 42 and issplit between the hydrogenation effluent recycle in the line 24 andolefin cracking feed in a line 44. The olefin cracking feed in the line44 is heated sufficiently to enter the vapor phase in the heater 32 andgaseous olefin cracking feed enters the olefin cracking reactor 34. Theheater 32 may comprise several stages of heating including a heatexchange with effluent from the olefin cracking reactor 34. In theolefin cracking reactor 34, the olefin cracking feed contacts olefincracking catalyst under gaseous phase conditions. Although the olefincracking reactor 34 is shown to be an upflow reactor, it may be orientedto be in a down flow or radial flow configuration. Upon contacting theolefin cracking catalyst under olefin cracking conditions, heavy olefinsin the C₄ to C₇ range crack down to light olefins in the C₂ to C₃ range.Cracked olefin effluent leaves the olefin cracking reactor 34 in a line46 and feeds a depropanizer column 50. An overhead C₃ ⁻ streamcomprising light olefins in a overhead line 52 is cooled in a condenser54 and split in a receiver 56 between a C₃ ⁻ product stream 58 and areflux stream 60. A bottoms stream in a line 62 is split between areboil line 64 and a debutanizer feed line 66. The reboil line 64 isheated in a reboiler 69 and returned to the depropanizer column 50 whilethe debutanizer feed line 66 is fed to the debutanizer column 16. Adebutanizer overhead line 68 carrying a C₄ ⁻ stream with olefins iscondensed in a condenser 70 and enters a receiver 72. The effluent fromthe receiver 72 in the line 22 is split between a debutanized processstream in the line 20, a purge stream in a line 76 and a reflux streamin a line 74. The bottoms stream in a line 78 comprising C₈ hydrocarbonsis split between a reboil stream 80 which is reboiled in a reboiler 82and returned to the debutanizer column 16 and a bottom purge portion ina line 84. An intermediate cut of C₅ to C₇ hydrocarbons with olefins ina line 14 from a side draw in the debutanizer column 16 is split betweena process side cut in the line 12 and a purge side cut in a line 86. Thepurges in the lines 76, 84 and 86 are all combined in a stream 90 to bemixed in a motor fuel pool or recycled upstream to a unit which providedfresh olefinic feed in the line 10. The purge provides for theelimination of C₄ ⁺ paraffins which would otherwise build up in theprocess.

FIG. 2 depicts an alternative embodiment of the present invention thatomits the need for one of the fractionation columns, preferably thedebutanizer column 16 in FIG. 1. A fresh olefinic feed comprising 20 to50 wt-% intermediate-weight C₄ to C₈ hydrocarbons with olefins entersthe process in a line 110. A line 112 carrying an intermediate cutcontaining intermediate-weight C₅ to C₇ hydrocarbons with olefins from adepropanizer column 150 mixes with a recycled hydrogenated effluent in aline 124 before mixing with the fresh olefinic feed in the line 110 toprovide an olefin hydrogenation feed in a line 128. The depropanizerfractionation column 150 includes several trays 102, but a tray 104 is atotal liquid accumulator that includes a liquid trap for collectingliquid and from which the line 112 takes a side draw, so that theintermediate cut in the line 112 is substantially all liquid. The olefinhydrogenation feed includes diolefins and acetylenes which can crack tomethane in an olefin cracking reactor 134. Hence, the olefinhydrogenation feed in the line 128 is admixed with hydrogen from a line136 and temperature controlled in a heat exchanger 138 before thehydrogen containing olefin hydrogenation feed in a line 140 enters aselective hydrogenation reactor 126. The olefin hydrogenation feed is inliquid phase and the hydrogen is dissolved into the liquidhydrocarbonaceous hydrogenation feed. The liquid olefin hydrogenationfeed enters the hydrogenation reactor 126 and contacts a preferablyfixed bed of hydrogenation catalyst. The hydrogenation reactor 126 maybe configured for up flow or radial flow, although down flow is shown inFIG. 2. Under selective hydrogenation conditions, diolefins andacetylenes in the olefin hydrogenation feed are converted to monoolefinswithout substantial monoolefin saturation. Hydrogenation effluent with asmaller concentration of diolefins than in the olefin hydrogenation feedin the line 128 exits the hydrogenation reactor 126 in a line 142 and issplit between the hydrogenation effluent recycle in the line 124 and thedepropanizer column feed in a line 143. The depropanizer column feed inthe line 143 is delivered to a depropanizer column 116 at a point thatis preferably below the point from which the liquid side draw is takenby the line 112. Lighter components of the hydrogenation effluent aredistilled upwardly in the depropanizer column 116 and a side vapor drawis taken preferably below the tray 104 from which the liquid draw istaken and above the point at which the line 143 delivers depropanizercolumn feed. The side vapor draw is taken through a line 144 andcomprises olefin cracking feed. The olefin cracking feed in the line 144is heated in a heat exchanger 132 and delivered to the olefin crackingreactor 134. Hence, the line 142 carrying hydrogenated effluent fluidlycommunicates with the line 144 carrying olefin cracking feed through thedepropanizer column 150. Because the olefin cracking feed in the line144 is a vapor draw from the depropanizer column 150, the heat exchanger132 does not have to provide heat of vaporization necessary to evaporatethe liquid hydrogenated effluent in the line 143, thereby requiring lessheat duty. In the olefin cracking reactor 134, the olefin cracking feedcontacts olefin cracking catalyst under gaseous phase conditions.Although the olefin cracking reactor 134 is shown to be an upflowreactor, it may be oriented to be in a down flow or radial flowconfiguration. Upon contacting the olefin cracking catalyst under olefincracking conditions, heavy olefins in the C₄ to C₇ range crack down tolight olefins in the C₂ to C₃ range. Cracked olefin effluent leaves theolefin cracking reactor 134 in a line 146 which feeds the depropanizercolumn 150, preferably at a point that is higher than the point at whichthe vapor side draw is taken by the line 144 and the point at which theliquid side draw is taken by the line 112. An overhead C₃ ⁻ streamcomprising light olefins in an overhead line 152 is cooled in acondenser 154 and split in a receiver 156 between a C₃ ⁻ product stream158 and a reflux stream 160. A bottoms stream in a line 162 comprisingC₈ ⁺ hydrocarbons is split between a reboil line 164 and a purge line166. The reboil line 164 is heated in a reboiler 168 and returned to thedepropanizer column 150 while the bottoms purge is removed in the purgeline 166.

Optionally, an intermediate cut of C₅ to C₇ hydrocarbons may be purgedby a purge line 186 at a point below the point at which vapor and liquidside draws are taken by the lines 144 and 112, respectively, and abovethe point at which hydrogenation effluent feeds the depropanizer column150 by the line 143. The liquid purge may be taken from a tray 106 bythe purge line 186 and mixed with the purge line 166 in a line 190. Anoptional condenser 188 either stabbed in or in fluid communication withthe depropanizer column 150 below the tray 104 and above the tray 106may be used to adjust the proportion of heavy paraffins in the feedwithdrawn in the line 144 and delivered to the olefin cracking reactor134. If the condenser 188 is used, a reflux line 192 may be used toreturn liquid to the depropanizer column 150. Optional trays 102′ areonly used if the optional equipment, the tray 106, the condenser 188,the purge line 186 and the reflux line 192, are utilized. Otherwise, thedepropanizer column 150 includes no trays between the tray 104 and thehydrogenation effluent inlet line 143.

1. An apparatus for the production of light olefins comprising: ahydrogenation reactor containing hydrogenation catalyst for selectivelyhydrogenating diolefins in an olefin stream with hydrogen to obtainmonoolefins, providing an olefin cracking feed; at least onehydrogenation effluent conduit fluidly communicating an outlet of saidhydrogenation reactor with an inlet to an olefin cracking reactor; meansfor vaporizing hydrogenated effluent in fluid communication with saidhydrogenation effluent conduit, wherein the means for vaporizing thehydrogenated effluent and the outlet of the hydrogenation reactor areconnected through a pipe without any intervening processing units; theolefin cracking reactor containing olefin cracking catalyst forconverting olefins in the olefin cracking feed into a cracked olefinstream comprising C₂ and C₃ olefins, and in communication with the meansfor vaporizing the hydrogenated effluent, wherein the olefin crackingreactor inlet and the means for vaporizing the hydrogenated effluent areconnected through a pipe without any intervening processing units; afractionation zone in fluid communication with said olefin crackingreactor for separating light hydrocarbons from heavy hydrocarbons insaid cracked olefin stream, said fractionation zone including at leastone fractionation column; and a side cut conduit fluidly communicating aside cut from said fractionation column to an olefin feed line formixing with the olefin feed to said hydrogenation reactor or to saidolefin cracking reactor.
 2. The apparatus of claim 1 wherein thefractionation zone includes two fractionation columns and the side cutis taken from the second fractionation column and the light olefinproduct is taken from the first fractionation column.
 3. The apparatusof claim 1 wherein the fractionation zone includes just one column. 4.The apparatus of claim 3 wherein the side cut is taken from a liquiddraw tray and the side cut conduit fluidly communicates with thehydrogenation reactor.
 5. The apparatus of claim 3 wherein the side cutis taken from a vapor stream and the side cut conduit fluidlycommunicates with said olefin cracking reactor.